Process for the conversion of ethane to aromatic hydrocarbons

ABSTRACT

A process for producing aromatic hydrocarbons which comprises (a) contacting ethane with a dehyroaromatization aromatic catalyst which is comprised of about 0.005 to about 0.1 wt % platinum, an amount of gallium which is equal to or greater than the amount of the platinum, from about 10 to about 99.9 wt % of an aluminosilicate, and a binder, and (b) separating methane, hydrogen, and C 2-5  hydrocarbons from the reaction products of step (a) to produce aromatic reaction products including benzene.

CROSS-REFERENCED TO RELATED APPLICATION

This application is a divisional application of U.S. Pat. No. 8,772,563,filed Feb. 16, 2009, which claims the benefit of U.S. ProvisionalApplication No. 61/029,478, filed on Feb. 18, 2008, which applicationsare hereby incorporated herein by reference.

FIELD OF THE INVENTION

The present invention relates to a process for producing aromatichydrocarbons from ethane. More specifically, the invention relates to adehydroaromatization process for increasing the production of benzeneand/or total aromatics from ethane.

BACKGROUND OF THE INVENTION

There is a projected global shortage for benzene which is needed in themanufacture of key petrochemicals such as styrene, phenol, nylon andpolyurethanes, among others. Generally, benzene and other aromatichydrocarbons are obtained by separating a feedstock fraction which isrich in aromatic compounds, such as reformates produced through acatalytic reforming process and pyrolysis gasolines produced through anaphtha cracking process, from non-aromatic hydrocarbons using a solventextraction process.

In an effort to meet growing world demand for benzene and otheraromatics, various industrial and academic researchers have been workingfor several decades to develop catalysts and processes to make lightaromatics (benzene, toluene, xylenes, or BTX) from cost-advantaged,light paraffin (C₁-C₄) feeds. Prior-art catalysts devised for thisapplication usually contain an acidic zeolite material such as ZSM-5 andone or more metals such as Pt, Ga, Zn, Mo, etc. to provide adehydrogenation function. For example, U.S. Pat. No. 7,186,871 describesa process for lower alkane aromatization utilizing a Pt/ZSM-5 catalyst.In this patent, data are presented for propane aromatization overvarious Pt/ZSM-5 catalysts, with the results showing better activity andselectivity to BTX as the Pt level of the catalyst is increased from0.06 to 0.33% wt. As another example, U.S. Pat. No. 4,350,835 covers aprocess for producing aromatic compounds from ethane-containing feedusing a Ga-containing zeolitic catalyst which may be of the ZSM-5 type.This latter patent teaches that the Ga content of the catalyst may varybetween 0.05 and 10% wt, preferably between 0.5 and 5% wt.

Aromatization of ethane and other lower alkanes is thermodynamicallyfavored at high temperature and low pressure without addition ofhydrogen to the feed. Unfortunately, these process conditions are alsofavorable for rapid catalyst deactivation due to formation ofundesirable surface coke deposits which block access to the activesites.

For many hydrocarbon processing applications, one approach to reducingcatalyst performance decline rates due to coking is to increase thecatalyst metals loading in an effort to promote fasterhydrogenation/breakup of large coke precursor molecules on the surface.Another approach involves incorporation of additives such as phosphateor rare earths to moderate surface acidity and reduce coking rates underreaction conditions. These approaches are appropriate for processesfeaturing fixed or slowly-moving catalyst beds wherein the averagecatalyst particle residence time in the reactor zone betweenregenerations (coke burnoff steps) is relatively long (at least severaldays). For example, see U.S. Pat. Nos. 4,855,522 and 5,026,937, whichdescribe ZSM-5-type lower-alkane aromatization catalysts promoted withGa and additionally containing either a rare earth metal or aphosporus-containing alumina, respectively.

Yet another approach to circumvent this problem is to devise a loweralkane aromatization process in which the catalyst spends a relativelyshort time (less than a day) under reaction conditions before beingsubjected to coke burnoff and/or other treatment(s) aimed at restoringall or some of the original catalytic activity. An example of such aprocess is one featuring two or more parallel reactors containing fixedor stationary catalyst beds, with at least one reactor offline forcatalyst regeneration at any given time, while the other reactor(s)is/are processing the lower alkane feed under aromatization conditionsto make aromatics. Another example of such a process features afluidized catalyst bed, in which catalyst particles cycle rapidly andcontinuously between a reaction zone where aromatization takes place anda regeneration zone where the accumulated coke is burned off thecatalyst to restore activity. For example, U.S. Pat. No. 5,053,570describes a fluid-bed process for converting lower paraffin mixtures toaromatics.

Requirements for optimal catalyst performance in a process involving arelatively short period of catalyst exposure to reaction conditionsbetween each regeneration treatment, such as a fluidized-bed process,can differ from those of fixed- or moving-bed processes which requirelonger catalyst exposure time to reaction conditions betweenregeneration treatments. Specifically, in processes involving shortcatalyst exposure times, it is important that the catalyst not exhibitexcessive initial cracking or hydrogenolysis activity which couldconvert too much of the feedstock to undesirable, less-valuablebyproducts such as methane.

Certain metals such as Pt which are very suitable for catalyzing thedehydrogenation reactions needed for aromatization can also, undercertain circumstances, display undesirable hydrogenolysis activity thatleads to excessive production of methane from higher hydrocarbons. Theinclusion of a second, inert or less-active metal in a catalystcomposition to help suppress the hydrogenolysis activity of the first,more-active metal is used in commercial scale catalytic naphthareforming in which C₅-C₁₂ paraffins and naphthenes are converted toaromatic compounds with catalysts which are predominantly bimetallic andare supported on chloride-promoted alumina. As indicated in a catalyticnaphtha reforming review article by C. A. Querini in volume 6, pages1-56 of the Encyclopedia of Catalysis (I. T. Horvath, ed.; published byJohn Wiley & Sons, Inc., Hoboken, N.J., USA, 2003), these catalyststypically contain Pt plus another metal such as Re (in sulfided form) orSn. Among other effects, these second metals can interact with the Pt toreduce hydrogenolysis activity, thereby decreasing the rate of unwantedmethane formation.

These Pt/Re and Pt/Sn catalysts, supported on chloride-promoted alumina,are widely employed in fixed-bed (semi-regenerative) and moving-bed(continuous) naphtha reformers, respectively, and their compositions areoptimized for relatively long catalyst exposure times to reactionconditions between regeneration treatments. The average catalystparticle residence time in the reaction zone between regenerationtreatments ranges from a few days in moving bed reactors and up to 1 or2 years in fixed bed reactors. According to the article by Querinimentioned above, typical Pt and Sn levels in Pt/Sn naphtha reformingcatalysts are about 0.3% wt each. Such catalysts, which usually lack astrongly acidic zeolite component, do not work well for lower alkanearomatization.

It would be advantageous to provide a light hydrocarbondehydroaromatization process which can be performed under conditionsthermodynamically favorable for light alkane aromatization as describedabove, which provides for relatively short catalyst exposure time toreaction conditions, wherein the average catalyst particle residencetime in the reaction zone between regeneration treatments may be fromabout 0.1 second to about 30 minutes in a fluidized bed reactor and froma few hours up to a week in moving bed and fixed bed reactors, and inwhich the catalyst composition is optimized to reduce excessive initialproduction of less-desirable byproducts such as methane.

SUMMARY OF THE INVENTION

The present invention provides a process for producing aromatichydrocarbons which comprises:

(a) contacting ethane with a dehydroaromatization catalyst wherein theethane contact time (the average residence time of a given ethanemolecule in the reaction zone under reaction conditions) is preferablyfrom about 0.1 seconds to about 1 minute, most preferably about 1 toabout 5 seconds, preferably at about 550 to about 730° C. and about 0.01to about 1.0 MPa, said catalyst comprising:

-   -   (1) about 0.005 to about 0.1 wt % (% by weight) platinum, basis        the metal, preferably about 0.01 to about 0.05% wt, most        preferably about 0.02 to about 0.05% wt,    -   (2) an amount of gallium which is equal to or greater than the        amount of the platinum, preferably no more than about 1 wt %,        most preferably no more than about 0.5 wt %;    -   (3) about 10 to about 99.9 wt % of an aluminosilicate,        preferably a zeolite, basis the aluminosilicate, preferably        about 30 to about 99.9 wt %, preferably selected from the group        consisting of ZSM-5, ZSM-11, ZSM-12, ZSM-23, or ZSM-35,        preferably converted to the H+ form, preferably having a        SiO₂/Al₂O₃ molar ratio of from about 20:1 to about 80:1, and    -   (4) a binder, preferably selected from silica, alumina and        mixtures thereof;

(b) collecting the products from (a) and separating and recovering C₆₊aromatic hydrocarbons;

(c) optionally recovering methane and hydrogen; and

(d) optionally recycling C₂₋₅ hydrocarbons to (a).

The reactor system may comprise one or more reaction vessels, chambers,or zones, arranged in parallel or in series, in which contact betweenthe catalyst particles and the ethane-containing feed occurs. Thereactor vessel(s), chamber(s), or zone(s) may feature a fixed catalystbed (i.e., with parallel beds), a slowly-moving catalyst bed, or afluidized bed, In a preferred embodiment, a fluidized-bed reactor isused. The process is optimized to minimize the average catalyst particleresidence time while maintaining selectivity and conversion rate. Theaverage catalyst particle residence time is the average amount of timethat a catalyst particle is in the active reaction zone with ethanebetween regenerations.

Catalysts of the present invention—featuring lower levels of a strongdehydrogenation metal (preferably Pt) with a potentially strong crackingactivity, complemented with gallium, which has lesser dehydrogenationactivity than Pt but can also suppress undesired cracking reactions—aredesigned to limit initial cracking activity without sacrificing theoverall activity and aromatics selectivity required forcommercially-viable production rates of benzene and other aromatics.

DETAILED DESCRIPTION OF THE INVENTION

The present invention is a process for producing aromatic hydrocarbonswhich comprises bringing a hydrocarbon feedstock containing at leastabout 50 percent by weight of ethane or other C₂ hydrocarbons intocontact with a dehydroaromatization catalyst composition suitable forpromoting the reaction of ethane to aromatic hydrocarbons such asbenzene at a temperature of about 550 to about 730° C. and a pressure ofabout 0.01 to about 1.0 MPa. The primary desired products of the processof this invention are benzene, toluene and xylene.

The hydrocarbons in the feedstock may be ethane, ethylene or mixturesthereof. Preferably, the majority of the feedstock is ethane and fromabout 0 to about 20 weight percent of the feedstock may be comprised ofethylene, preferably about 5 to about 10 weight percent. The feedstockmay contain in addition up to about 40 weight percent of other openchain hydrocarbons containing between 3 and 8 carbon atoms ascoreactants. Specific examples of such additional coreactants arepropane, propylene, n-butane, isobutane, n-butenes and isobutene. Thehydrocarbon feedstock preferably contains at least about 60 percent byweight of C₂ hydrocarbons, more preferably at least about 70 percent byweight. The reaction feed is often referred to herein as ethane forconvenience but it is meant to include all of the other hydrocarbonmaterials referred to above if it is necessary or desired for them to bepresent.

In a preferred embodiment, the reactor comprises a zone, vessel, orchamber containing catalyst particles through which theethane-containing feed flows and the reaction takes place. The reactorsystem may involve a fixed, moving, or fluidized catalyst bed. Thereaction products then flow out of the bed and are collected. Thereaction products are then separated and C₆₊ aromatic hydrocarbons arerecovered. Optionally, methane and hydrogen are recovered and optionallythe C₂₋₅ hydrocarbons are recycled to step (a).

A fixed bed reactor is a reactor in which the catalyst remainsstationary in the reactor and the catalyst particles are arranged in avessel, generally a vertical cylinder, with the reactants and productspassing through the stationary bed. In a fixed bed reactor the catalystparticles are held in place and do not move with respect to a fixedreference frame. The fixed bed reactor may be an adiabatic single bed, amulti-tube surrounded with heat exchange fluid or an adiabatic multi-bedwith internal heat exchange, among others. Fixed bed reactors are alsoreferred to as packed bed reactors. Fixed bed reactors provide excellentgas solids contacting. The fixed bed reactor configuration may includeat least two separate fixed beds in different zones so that at least onebed can be in active operation under reaction conditions while thecatalyst the other bed(s) is being regenerated.

In a moving bed catalytic reactor, gravity causes the catalyst particlesto flow while maintaining their relative positions to one another. Thebed moves with respect to the wall of the vessel in which it iscontained. The reactants may move through this bed with cocurrent,countercurrent or crossflow. Plug flow is the preferred mode. The movingbed offers the ability to withdraw catalyst particles continuously orintermittently so they can be regenerated outside the reactor andreintroduced into the circuit later on. Thus, there is an advantage tousing a moving bed when the catalyst has a short active life and can becontinuously regenerated. A moving bed reactor may consist of at leastone tray as well as supporting means for one or more catalyst beds. Thesupporting means may be permeable to gas and impermeable to catalystparticles.

A fluidized bed reactor is a type of reactor that may be used to carryout a variety of multiphase chemical reactions. In this type of areactor, a gas is passed through the particulate catalyst at high enoughvelocities to suspend the solid and cause it to behave as though it werea fluid. The catalyst particles may be supported by a porous plate. Thegas may be forced through the porous plate up through the solidmaterial. At lower gas velocities the solids remain in place as the gaspasses through the voids in the material. As the gas velocity isincreased, the reactor reaches the stage where the force of the fluid onthe solids is enough to balance the weight of the solid material andabove this velocity the contents of the reactor bed begin to expand andswirl around much like an agitated tank or boiling pot of water. Afluidized bed reactor is preferred for use in the present inventionbecause it provides uniform particle mixing, uniform temperaturegradients and the ability to operate the reactor in a continuous state.The catalyst leaves the reaction zone with the reaction products and isseparated therefrom in order to be regenerated before being recycled tothe reaction zone.

The ethane contact time may range from about 0.1 second to about 1minute. The ethane contact time is the average amount of time that onemolecule of the ethane feed is in the reaction zone. The preferredethane contact time is from about 1 to about 5 seconds. Longer ethanecontact times are less desirable because they tend to allow forsecondary reactions that lead to less-desirable byproducts such asmethane and reduce selectivity to benzene and/or total aromatics.

The catalyst comprises from about 0.005 to about 0.1 wt % platinum,basis the metal. The platinum is highly active in terms of catalyzingthe dehydroaromatization reaction and it is best if its concentration inthe catalyst not be more than 0.1 wt % because otherwise too muchmethane will be produced. In one embodiment from about 0.01 to about0.05 wt % platinum is used, preferably from about 0.02 to about 0.05 wt% of platinum is used. High performance is thus obtained with relativelylow amounts of metals in the catalyst.

A second metal, namely gallium (Ga), is an essential component of thecatalyst of the present invention. The Ga helps to suppress undesirablecatalytic activity which leads to the production of less-valuablemethane byproduct, leading to an improvement in selectivity to benzeneand total aromatics without excessively dampening the total level ofethane conversion or aromatics production. For the present invention,the amount of Ga, by weight, in the catalyst composition may be equal toor greater than the amount of Pt, basis the metal. In one embodiment nomore than about 1 wt % of gallium may be used. Preferably no more thanabout 0.5 wt %, most preferably no more than about 0.3 wt %, is used.

The catalyst also comprises from about 10 to about 99.9 wt % of one ormore aluminosilicate materials, preferably from about 30 to about 99.9wt %, basis the aluminosilicate(s). The aluminosilicates preferably havea silicon dioxide:aluminum trioxide molar ratio of from about 20 toabout 80. The aluminosilicates may preferably be zeolites having the MFIor MEL type structure and may be ZSM-5, ZSM-8, ZSM-11, ZSM-12 or ZSM-35.The zeolite or zeolite mixture is preferably converted to H⁺ form toprovide sufficient acidity to help catalyze the dehydroaromatizationreaction. This can be accomplished by calcining the ammonium form of thezeolite in air at a temperature of at least about 400° C.

The binder material serves the purpose of holding invidivual zeolitecrystal particles together to maintain an overall catalyst particle sizein the optimal range for fluidized-bed operation or to prevent excessivepressure drop in fixed or moving bed operation. The binder may beselected from the group consisting of alumina, silica, silica/alumina,various clay materials such as kaolin, or mixtures thereof. Preferably,amorphous inorganic oxides of gamma alumina, silica, silica/alumina or amixture thereof may be included. Most preferably, alumina and/or silicaare used as the binder material.

A platinum containing crystalline aluminosilicate, such as ZSM-5, may besynthesized by preparing the aluminosilicate containing the aluminum andsilicon in the framework, depositing platinum on the aluminosilicate andthen calcining the aluminosilicate. The gallium may also be added by thesame procedure, either prior to, simultaneously with, or after theaddition of platinum. The metals may be added by any commonly knownmethod for adding metals to such structures including incorporation intothe aluminosilicate framework during crystal synthesis, or subsequention exchange into an already-synthesized crystal framework, or well asby various impregnation methods known to those skilled in the art. Theplatinum and gallium may be added by the same or different methods.

In a preferred embodiment of the present invention an ethane feedstreamis introduced into the dehydroaromatization reactor. The feedstream thencomes into contact with the catalyst particles for the prescribed periodof time. The reaction products leave the reactor and are transferredinto a separator. The separator removes the aromatic products and theprincipal byproducts, methane and hydrogen, which preferably may berecovered, and also removes C₂₋₅ byproducts and unreacted ethane whichoptionally may be recycled to the dehydroaromatization reactor.

EXAMPLES

The following example is illustrative only and is not intended to limitthe scope of the invention.

Example 1

Catalysts A through H were prepared on samples of an extrudate materialcontaining 80% wt of CBV 3014E ZSM-5 zeolite (30:1 molar SiO₂:Al₂O₃ratio, available from Zeolyst International) and 20% wt of aluminabinder. The extrudate samples were calcined in air up to 425° C. toremove residual moisture prior to use in catalyst preparation.

Metals were deposited on 25-50 gram samples of the above ZSM-5/aluminaextrudate by first combining appropriate amounts of stock solutions oftetraammine platinum nitrate and gallium(III) nitrate, diluting thismixture with deionized water to a volume just sufficient to fill thepores of the extrudate, and impregnating the extrudate with thissolution at room temperature and atmospheric pressure. Impregnatedsamples were aged at room temperature for 2-3 hours and then driedovernight at 100° C.

To determine the platinum and gallium contents of each catalyst, asample of the catalyst was calcined at 550° C. to drive off residualmoisture to render a loss on ignition (LOI) percentage. A known mass ofthe untreated ground catalyst, corrected by LOI percentage, was digestedusing closed vessel microwave acid digestion involving nitric,hydrochloric, and hydrofluoric acids. The solution was diluted to aknown volume with deionized water and then analyzed for the indicatedmetals by directly coupled plasma emission analysis. Results arereported as ppmw or weight percent based on the weight of the 550°C.-calcined catalyst sample.

Catalysts made on the ZSM-5/alumina extrudate were tested “as is,”without crushing. For each performance test, a 15-cc charge of catalystwas loaded into a quartz tube (1.40 cm i.d.) and positioned in athree-zone furnace connected to an automated gas flow system.

Prior to performance testing, all catalyst charges were pretreated insitu at atmospheric pressure as follows:

-   -   (a) calcination with air at 60 liters per hour (L/hr), during        which the reactor wall temperature was increased from 25 to        510° C. in 12 hrs, held at 510° C. for 4-8 hrs, then further        increased from 510 to 630° C. in 1 hr, then held at 630° C. for        30 min;    -   (b) nitrogen purge at 60 L/hr, 630° C. for 20 min;    -   (c) reduction with hydrogen at 60 L/hr, 630° C. for 30 min.

At the end of the pretreatment, 100% ethane feed was introduced at 15L/hr (1000 gas hourly space velocity-GHSV), atmospheric pressure, withthe reactor wall temperature maintained at 630° C. The total reactoroutlet stream was sampled and analyzed by an online gas chromatographysystem two minutes after ethane feed addition. Based on composition dataobtained from the gas chromatographic analysis, initial ethaneconversion and hydrocarbon product selectivities were computed accordingto the following formulas:ethane conversion, %=100×(100−% wt ethane in outlet stream)/(% wt ethanein feed)selectivity to hydrocarbon product Y (other than ethane)=100×(moles ofcarbon in amount of product Y generated)/(moles of carbon in amount ofethane reacted)For purposes of the selectivity calculation, C₉₊ aromatics were assumedto have an average molecular formula of C₁₀H₈ (naphthalene).

The analyzed platinum and gallium levels and initial aromatizationperformance data for Catalysts A-H, prepared and tested as describedabove, are presented in Table 1. The data in Table 1 indicate that thelow-Pt/Ga/ZSM-5 catalysts B through F of the present invention providebetter initial suppression of methane production and higher selectivityto benzene and total aromatics under ethane aromatization conditionsthan catalysts A, G, and H, in which the Pt and/or Ga levels falloutside the ranges of the present invention.

TABLE 1 Catalyst A B C D E F G H Analyzed Pt Level, % wt 0.018 0.0240.022 0.024 0.0275 0.028 0.198 0.213 Analyzed Ga Level, % wt 0.014 0.0250.092 0.150 0.472 0.912 0.145 0.917 Ethane conversion, % 44.88 43.9543.5 46.49 48.25 46.93 63.24 50.04 Selectivities, % (carbon basis)Methane 26.45 18.43 16.95 17.78 16.08 18.1 42.65 22.50 Ethylene 12.7912.25 10.15 12.30 10.82 11.16 8.2 9.98 Propylene 1.84 1.74 1.48 1.571.39 1.44 0.64 1.14 Propane 1.65 1.85 1.97 1.57 1.46 1.54 0.5 1.24 C4Hydrocarbons 0.36 0.37 0.35 0.35 0.34 0.30 0.16 0.26 C5 Hydrocarbons0.01 0.03 0.04 0.03 0.01 0.01 0.01 0.01 Benzene 34.21 38.13 37.09 38.5436.02 36.76 30.41 36.47 Toluene 17.64 19.74 19.89 19.43 17.86 17.8213.11 16.93 C8 Aromatics 3.16 3.42 4.05 3.33 3.37 2.76 2.03 2.55 C9+Aromatics 1.90 4.04 8.05 5.09 12.64 10.11 2.3 8.94 Total Aromatics 56.9165.33 69.07 66.39 69.89 67.45 47.85 64.88

What is claimed is:
 1. A catalyst comprising: (a) about 0.01 to about0.05 wt% platinum, based on the total weight of the catalyst calculatedas metal, (b) an amount of gallium equal to or greater, by weight, thanthe amount of the platinum, and no more than about 1 wt%, based on thetotal weight of the catalyst calculated as metal; (c) about 10 to about99.9 wt% of an aluminosilicate, wherein the aluminosilicate is a zeolitehaving an MFI or MEL type structure, and (d) a binder.
 2. The catalystof claim 1 wherein the catalyst comprises an amount of gallium of nomore than about 0.5 wt%.
 3. The catalyst of claim 1 wherein the amountof the aluminosilicate is from about 30 to about 99.9 wt%.
 4. Thecatalyst of claim 1 wherein the aluminosilicate has a silicondioxide:aluminum trioxide molar ratio of from about 20 to about 80.